Integrated process for optimum production of para-xylene

ABSTRACT

A method of producing p-xylene comprising the steps of separating the reformate feed in the reformate splitter to produce a benzene stream, a combined heavy stream, a xylene stream, and a toluene stream, converting the C9+ aromatic hydrocarbons in the presence of a dealkylation catalyst in the dealkylation reactor to produce a dealkylation effluent, separating the dealkylation effluent in the dealkylation splitter to produce a C9 stream and a C10+ stream, reacting the C9 stream, the toluene stream, the benzene stream, and the hydrogen stream in the presence of a transalkylation catalyst in the transalkylation reactor to produce a transalkylation effluent, separating the p-xylenes from the xylene stream in the p-xylene separation unit to produce a p-xylene product and a p-xylene depleted stream, converting the m-xylene and o-xylene in the p-xylene depleted stream in the isomerization unit to produce an isomerization effluent.

REFERENCE TO CROSS-RELATED APPLICATIONS

This application is divisional application of U.S. Non-Provisionalpatent application Ser. No. 16/592,158 filed on Oct. 3, 2019, which is acontinuation-in-part application of U.S. Non-Provisional patentapplication Ser. No. 16/160,393 filed on Oct. 15, 2018 and issued asU.S. Pat. No. 10,696,609 on Jun. 30, 2020. For purposes of United Statespatent practice, this application incorporates the contents of bothNon-Provisional Application by reference in their entirety.

TECHNICAL FIELD

Disclosed are methods and systems for production of xylenes.Specifically, disclosed are methods and systems for maximizingproduction of xylenes from reformate.

BACKGROUND

In a typical aromatics complex catalytic naphtha reformate is directedto benzene-toluene-xylene (BTX) separation steps and the remaining C9+fraction can be separated into a C9 fraction and a C10+ fraction. TheC10+ fraction can be rejected from the system.

The C9 fraction can be introduced to a transalkylation reactor toproduce additional BTX. The extracted xylenes can be directed to apara-xylene separation step. The remaining xylene isomers can be sent toa stand-alone xylene isomerization reactor to re-establish thethermodynamic equilibrium of xylenes and form para-xylene in theprocess.

Benzene can be separated from the catalytic naphtha reformate via theextraction process. In addition, benzene can be produced in the toluenedisproportionation reactor, through thermal dealkylation, orhydrodealkylation. The produced benzene is recovered from the process.In a typical process, the benzenes recovered or produced are notrecycled within the process. Additional xylenes can be produced in thethermal dealkylation and hydrodealkylation.

SUMMARY

Disclosed are methods and systems for production of xylenes.Specifically, disclosed are methods and systems for maximizingproduction of xylenes from reformate.

In a first aspect, a method of producing para-xylene (p-xylene) isprovided. The method includes the steps of introducing a reformate feedto reformate splitter, the reformate splitter configured to separate thereformate feed, the reformate feed includes aromatic hydrocarbons, wherethe aromatic hydrocarbons are selected from the group consisting ofbenzene, toluene, mixed xylenes, carbon-nine plus (C9+) aromatichydrocarbons, and combinations of the same and separating the reformatefeed in the reformate splitter to produce a light gases stream, abenzene stream, a combined heavy stream, a xylene stream, and a toluenestream, where the combined heavy stream includes C9+ aromatichydrocarbons, and where the xylene stream includes mixed xylenes, wherethe mixed xylenes includes p-xylenes, and where the toluene streamincludes toluene. The method further includes the steps of introducingthe combined heavy stream to a dealkylation reactor, introducing ahydrogen feed to the dealkylation reactor, where the hydrogen feedincludes hydrogen gas, where the dealkylation reactor is configured tothe convert C9+ aromatic hydrocarbons to carbon-six (C6) to carbon-eight(C8) aromatic hydrocarbons, and converting the C9+ aromatic hydrocarbonsand the hydrogen gas in the presence of a dealkylation catalyst in thedealkylation reactor to produce a dealkylation effluent, where thedealkylation reactor is at a dealkylation temperature, where thedealkylation reactor is at a dealkylation pressure, where thedealkylation effluent includes aromatic hydrocarbons such that an amountof C9+ aromatic hydrocarbons in the dealkylation effluent is less thanthe amount of C9+ aromatic hydrocarbons in the reformate feed. Themethod further includes the steps of introducing the dealkylationeffluent to a dealkylation splitter, the dealkylation splitterconfigured to separate the dealkylation effluent, separating thedealkylation effluent in the dealkylation splitter to produce a C9stream, a C10+ stream, and a light recycle stream, where the C10+ streamincludes C10+ aromatic hydrocarbons, where the C9 stream includes C9aromatic hydrocarbons, introducing the C9 stream and a hydrogen streamto a transalkylation reactor, the transalkylation reactor configured toconvert C9 aromatic hydrocarbons through transalkylation reactions,where the hydrogen stream includes hydrogen gas, introducing the benzenestream to the transalkylation reactor, introducing the toluene stream asa toluene split stream to the transalkylation reactor, where a flow rateof the toluene split stream is operable to maintain a ratio of tolueneto trimethylbenzene in the transalkylation reactor in the range of 0.3to 3, and reacting the C9 stream, the toluene stream, the benzenestream, and the hydrogen stream in the presence of a transalkylationcatalyst in the transalkylation reactor to produce a transalkylationeffluent, where the transalkylation reactor is at a transalkylationtemperature, where the transalkylation reactor is at a transalkylationpressure, where the transalkylation catalyst is operable to catalyzetransalkylation reactions, where the transalkylation effluent includesC6 to C9+ aromatic hydrocarbons. The method further includes the stepsof introducing the xylene stream to a p-xylene separation unit, thep-xylene separation unit configured to separate p-xylenes from thexylene stream, separating the p-xylenes from the xylene stream in thep-xylene separation unit to produce a p-xylene product and a p-xylenedepleted stream, where the p-xylene product includes p-xylenes, wherethe p-xylene depleted stream includes m-xylene and o-xylene, introducingthe p-xylene depleted stream to an isomerization unit, where theisomerization unit includes an isomerization catalyst, and convertingthe m-xylene and o-xylene in the p-xylene depleted stream in theisomerization unit to produce an isomerization effluent, where theisomerization unit is at an isomerization temperature, where theisomerization unit is at an isomerization pressure, where theisomerization effluent includes C8 aromatic hydrocarbons. The methodfurther includes the step of introducing the light recycle stream, theisomerization effluent and the transalkylation effluent to the reformatesplitter.

In certain aspects, the method further includes the steps of introducinga fraction of the toluene stream not in the toluene split stream to atoluene disproportionation reactor, where the toluene disproportionationreactor includes a disproportion catalyst, where the toluenedisproportionation reactor is configured to support a toluenedisproportionation reaction, reacting the fraction of the toluene streamin the toluene disproportionation reactor to produce adisproportionation effluent, where the disproportionation effluentincludes C6 to C8 aromatic hydrocarbons, and introducing thedisproportionation effluent to the reformate column. In certain aspects,the method further includes the step of separating a fraction of theC10+ stream in a purge stream. In certain aspects, the dealkylationtemperature is between 200 deg C and 550 deg C, further where thedealkylation pressure is between 5 bar and 50 bar, and further where aliquid hourly space velocity in the dealkylation reactor is between 1hr-1 and 20 hr-1. In certain aspects, the transalkylation temperature isbetween 200 deg C and 550 deg C, further where the transalkylationpressure is between 10 bar and 50 bar, further where a liquid hourlyspace velocity in the transalkylation reactor is between 0.2 hr-1 and 20hr-1. In certain aspects, the isomerization temperature is between 200deg C and 550 deg C, further where the isomerization pressure is between5 bar and 50 bar, and further where a liquid hourly space velocity inthe isomerization unit is between 1 hr-1 and 20 hr-1.

In a second aspect, a system for producing para-xylene (p-xylene) from areformate feed is provided. The system includes a reformate splitter,the reformate splitter configured to separate the reformate feed, wherethe reformate feed includes aromatic hydrocarbons, where the aromatichydrocarbons are selected from the group consisting of benzene, toluene,mixed xylenes, carbon-nine plus (C9+) aromatic hydrocarbons, andcombinations of the same, where the reformate splitter produces a lightgases stream, a combined heavy stream, a benzene stream, a xylenestream, and toluene stream, where the C9 aromatics stream includes C9aromatic hydrocarbons, and where the xylene stream includes mixedxylenes, where the mixed xylenes includes p-xylenes, and where the mixedlight aromatics stream includes toluene. The system further includes adealkylation reactor fluidly connected to the reformate splitter, thedealkylation reactor configured to convert C9+ aromatic hydrocarbons tocarbon-six (C6) to carbon-eight (C8) aromatic hydrocarbons in thepresence of a dealkylation catalyst to produce a dealkylation effluent,where the dealkylation reactor is at a dealkylation temperature, wherethe dealkylation reactor is at a dealkylation pressure, where thedealkylation effluent includes aromatic hydrocarbons such that an amountof C9+ aromatic hydrocarbons in the dealkylation effluent is less thanthe amount of C9+ aromatic hydrocarbons in the reformate feed and adealkylation splitter fluidly connected to the dealkylation reactor, thedealkylation splitter configured to separate the dealkylation effluentto produce a light recycle stream, C10+ aromatics stream, and C9 stream.The system further includes a transalkylation reactor fluidly connectedto the dealkylation reactor and fluidly connected to the reformatesplitter, the transalkylation reactor configured to convert toluene andC9 aromatic hydrocarbons in the presence of a transalkylation catalystto produce a transalkylation effluent, where the transalkylation reactoris at a transalkylation temperature, where the transalkylation reactoris at a transalkylation pressure, where the toluene is supplied by atoluene split stream, where the toluene split stream is separated fromthe toluene stream, where the transalkylation catalyst is operable tocatalyze transalkylation reactions, where the transalkylation effluentincludes C6 to C9+ aromatic hydrocarbons, a p-xylene separation unitfluidly connected to a reformate splitter, the p-xylene separation unitconfigured to separate p-xylenes from the xylene stream to produce ap-xylene product and a p-xylene depleted stream, where the p-xyleneproduct includes p-xylenes, where the p-xylene depleted stream includesm-xylene and o-xylene, and an isomerization unit fluidly connected tothe p-xylene separation unit, where the isomerization unit is configuredto convert the m-xylene and o-xylene in the p-xylene depleted stream toproduce an isomerization effluent, where the isomerization unit is at anisomerization temperature, where the isomerization unit is at anisomerization pressure, where the isomerization effluent includes C8aromatic hydrocarbons.

In certain aspects, a flow rate of the toluene split stream is operableto maintain a ratio of toluene to trimethylbenzene in thetransalkylation reactor in a range of 0.3 to 3. In certain aspects, thesystem further includes a toluene disproportionation reactor fluidlyconnected to the reformate splitter, the toluene disproportionationreactor configured to convert a fraction of the toluene from the toluenestream to produce a disproportionation effluent, where thedisproportionation effluent includes C6 to C8 aromatic hydrocarbons.

BRIEF DESCRIPTION OF THE DRAWINGS

These and other features, aspects, and advantages of the scope willbecome better understood with regard to the following descriptions,claims, and accompanying drawings. It is to be noted, however, that thedrawings illustrate only several embodiments and are therefore not to beconsidered limiting of the scope as it can admit to other equallyeffective embodiments.

FIG. 1 provides a process diagram of an embodiment of the process.

FIG. 2 provides a process diagram of an embodiment of the process.

FIG. 3 provides a process diagram of an embodiment of the process.

FIG. 4 provides a process diagram of an embodiment of the process.

FIG. 5 provides a process diagram of the process for Example 2.

FIG. 6 provides a process diagram of an embodiment of the process.

FIG. 7 provides a process diagram of the process for Example 4.

In the accompanying Figures, similar components or features, or both,may have a similar reference label.

DETAILED DESCRIPTION

While the scope of the apparatus and method will be described withseveral embodiments, it is understood that one of ordinary skill in therelevant art will appreciate that many examples, variations andalterations to the apparatus and methods described here are within thescope and spirit of the embodiments.

Accordingly, the embodiments described are set forth without any loss ofgenerality, and without imposing limitations, on the embodiments. Thoseof skill in the art understand that the scope includes all possiblecombinations and uses of particular features described in thespecification.

Described here are processes and systems of a system to maximizeproduction of mixed xylenes. A reformate is introduced to a dealkylationreactor. The dealkylation effluent from the dealkylation reactor isseparated into the separate components in a reformate splitter. Themixed xylene stream is then separated to produce a p-xylene productstream. The p-xylene depleted stream is introduced to an isomerizationunit to produce a mixed xylene stream and the mixed xylene stream isthen introduced to a splitter column. The C9 aromatic hydrocarbons andtoluene are introduced to a transalkylation reactor along with a benzenerecycle, toluene recycle, and C9+ aromatic recycle from the splittercolumn. The transalkylation effluent is introduced to the splittercolumn where the effluent is separated into it components, thusproducing the benzene recycle, the toluene recycle, and the C9+ aromatichydrocarbons. The mixed xylenes separated from transalkylation effluentin the splitter column are recycled to the p-xylene separation.

The use of a dealkylation reactor can reduce the flow rate of heavyreformate going to the transalkylation reactor due to the conversion ofmethylethylbenzene to toluene in the dealkylation reactor.Advantageously, the use of a dealkylation reactor increases theproduction of mixed xylenes due to an increase in the conversion of C10+aromatic hydrocarbons in the dealkylation reactor and a reduced amountor absence of C10+ aromatic hydrocarbons in the transalkylation reactor.Advantageously, the process to maximize the production of xylenesutilizes the entire reformate stream to produce valuable aromatics, byintroducing the entire reformate stream to the dealkylation reactorbefore separating the stream to reject the C10+ fraction. In at leastone embodiment, the C10+ aromatic hydrocarbons in the dealkylationreactor effluent can be inert and therefore do not contribute to theproduction of xylenes. Examples of inert C10+ aromatic hydrocarbons inthe dealkylation reactor effluent include methyl-naphthalene (a C11aromatic) and diethylbenzene (a C10 aromatic). Such inert C10+ aromatichydrocarbons can cause deactivation of the catalyst in thetransalkylation reactor and effectively lower the efficiency oftransalkylation catalyst by diluting the feed. Advantageously, the useof the reformate splitter and the splitter column to separate thestreams into component parts enhances the ability to process thecomponents and increase the production of mixed xylenes as compared tobroader cut separations. Advantageously the combination of the reformatesplitter and the splitter column contributes to the increased productionof mixed xylenes due to the ability to direct and recycle the use ofeach component stream. Advantageously, the combination of a dealkylationreactor and a separate transalkylation reactor increases the overallproduction of xylenes as compared to a one-reactor system that containsboth a dealkylation catalyst and a transalkylation catalyst or a singlecatalyst capable of both dealkylation and transalkylation reactions.Advantageously, the use of separate units to perform disproportionationreactions, isomerization reactions, and dealkylation reactions, canallow each unit to be designed for a specific feed and specific catalystand can therefore be optimized to maximize production of the targetcomponent.

As used throughout, a reference to “C” and a number or “carbon-number”refers to the number of carbon atoms in a hydrocarbon. For example, C 1refers to a hydrocarbon with one carbon atom and C6 refers to ahydrocarbon with six carbon atoms.

As used throughout, “C9 aromatic hydrocarbons” refers to aromatichydrocarbons with nine carbon atoms. Examples of C9 aromatichydrocarbons include methylethylbenzene, trimethylbenzene,propylbenzene, and combinations of the same.

As used throughout, “trimethylbenzene” or “TMB” refers to and includeseach of the isomers of trimethylbenzene: hemellitene, pseudocumene, andmesitylene.

As used throughout, “C10+ aromatic hydrocarbons” refers to aromatichydrocarbons with ten carbon atoms and aromatic hydrocarbons with morethan ten carbon atoms, such as an aromatic hydrocarbon with elevencarbon atoms.

As used throughout, “C9+ aromatic hydrocarbons” refers to the group ofC9 aromatic hydrocarbons and C10+ aromatic hydrocarbons.

As used throughout, “mixed xylenes” refers to one or more of para-xylene(p-xylene), meta-xylene (m-xylene), and ortho-xylene (o-xylene).

As used throughout, “dealkylation reaction” refers to a reaction thatresults in the removal of an alkyl group from one or more of thereactants. For example, a reaction to convert methylethylbenzene andhydrogen to ethane and p-xylene is a dealkylation reaction.

As used throughout, “transalkylation reaction” refers to a reaction thatresults in the transfer of an alkyl group from one or compound toanother.

As used throughout, “isomerization reaction” refers to a reaction thatresults in the rearrangement of molecules to a different molecularstructure, but the molecular formula remains the same. For example, areaction to convert o-xylene to p-xylene or ethylbenzene to p-xylene isan isomerization reaction.

As used throughout, “light gases” refers to light hydrocarbons,hydrogen, air, and combinations of the same.

As used throughout, “light hydrocarbons” refers to alkanes, includingmethane, ethane, propane, butanes, pentanes, alkenes, and trace amountsof naphthenes, such as cyclopentane, cyclohexane, and combinations ofthe same.

Referring to FIG. 1 an embodiment of the process for increasing xyleneproduction is provided.

Reformate feed 2 and hydrogen feed 4 are introduced to dealkylationreactor 100. Reformate feed 2 can include C6 to C12 one ring aromatichydrocarbons. Reformate feed 2 can be sourced from a reformer unit. Areformer unit converts naphthas into aromatics-rich products and can useplatinum-containing catalyst at high temperatures and hydrogen pressuresto effect such conversions. The goal of a reformer unit is to increasethe octane number of the reformer feed stream. The primary reactions ina reformer unit include dehydrogenation reactions, however, crackingreactions can also occur.

In at least one embodiment, reformate feed 2 can include benzene,toluene, mixed xylenes, C9 aromatic hydrocarbons, C10+ aromatichydrocarbons, and combinations of the same. In at least one embodiment,the C10+ aromatic hydrocarbons can include C10 to C12 hydrocarbons.Reformate feed 2 can include between 65 percent by weight (wt %) and 80wt %. Reformate feed 2 can further include paraffins and naphthenes inthe range between 20 wt % and 35 wt %. The amount of aromatics andnon-aromatics depends on the nature of the feed to the reformate unitand the severity of the conditions in the reactors of the reformateunit.

Hydrogen feed 4 can be any stream containing hydrogen gas. Hydrogen feed4 can be a stream of pure hydrogen from a fresh hydrogen source. In atleast one embodiment, hydrogen gas can be recovered from another part ofthe process and can be recycled as hydrogen feed 4 and introduced todealkylation reactor 100. Hydrogen feed 4 can be from a hydrogen sourcein a refinery and can contain light hydrocarbons. The hydrogen gas inhydrogen feed 4 can ensure that the non-aromatics are not cracked toproduce olefins. One of skill in the art understands that olefins arenot formed in the presence of hydrogen, but paraffins can be. In analternate embodiment for increasing xylene production, dealkylationreactor 100 can be in the absence of hydrogen, such as in a fluidizedbed reactor, and olefins can be produced.

Dealkylation reactor 100 can be any type of reactor capable ofcontaining and supporting a dealkylation reaction. Dealkylation reactor100 can be a fixed bed reactor or a fluidized bed reactor. Inembodiments with hydrogen feed 4, dealkylation reactor 100 can be afixed bed reactor. In embodiments where dealkylation reactor 100 is inthe absence of hydrogen, dealkylation reactor 100 can be a fluidized bedreactor. The dealkylation temperature in dealkylation reactor 100 can bebetween 200 degrees Celsius (deg C) and 550 deg C. The dealkylationpressure in dealkylation reactor 100 can be between 5 bar (500kilopascal (kPa)) and 50 bar (5000 kPa) and alternately between 10 bar(1000 kPa) and 50 bar (5000 kPa). The liquid hourly space velocity(LHSV) can be between 1 per hour (hr⁻¹) and 20 hr⁻¹ and alternatelybetween 1 hr⁻¹ and 10 hr⁻¹. The volumetric flow ratio of hydrogen feed 4to reformate feed 2 can be in the range of 0 to 8.

Dealkylation reactor 100 can include a dealkylation catalyst. Thedealkylation catalyst can include any catalysts capable of catalyzingdealkylation reactions. Examples of dealkylation catalyst can includecatalysts such as those described in U.S. Pat. No. 6,096,938 and thosedescribed in U.S. Pat. No. 9,000,247. The dealkylation catalyst can beselected to selectively convert one or more of the C9+ aromatichydrocarbons over the others in dealkylation reactions. Dealkylationreactions can convert C9+ aromatic hydrocarbons to C6 to C8 aromatichydrocarbons. Dealkylation reactions can convert C10+ aromatichydrocarbons to C9 aromatic hydrocarbons. Reactions in dealkylationreactor 100 can remove methyl, ethyl, propyl, butyl and pentyl groups,and their isomers, attached to C10+ aromatic hydrocarbons. In at leastone embodiment, a dealkylation catalyst can be selected to convert 98 wt% of the methylethylbenzene in reformate feed 2 to toluene, andalternately greater than 98 wt %. In at least one embodiment, theoverall conversion of C9+ aromatic hydrocarbons to C6 to C8 aromaticscan be greater than 98 wt % due to conversion of C9 aromatichydrocarbons and the removal of methyl, ethyl, propyl, butyl and pentylgroups attached to C10+ aromatic hydrocarbons. Dealkylation effluent 6can contain mixed xylenes, toluene, benzene, light gases, and C9+aromatic hydrocarbons. Dealkylation effluent 6 can contain a greateramount of C6 to C8 aromatic hydrocarbons and a lesser amount of C9+aromatic hydrocarbons compared to reformate feed 2. Dealkylationeffluent 6 can contain a greater amount of p-xylene than reformate feed2.

Dealkylation effluent 6 is introduced to reformate splitter 200.

Reformate splitter 200 can be any type of separation unit capable ofseparating a stream into its component parts. In at least oneembodiment, reformate splitter 200 can be one splitter column designedto separate the feed stream into multiple split streams. In at least oneembodiment, reformate splitter 200 can be multiple splitter columns inseries designed to separate one component from the feed stream. In atleast one embodiment, reformate splitter 200 can be one or moredistillation units. The pressure across reformate splitter 200 can be inthe range from 0.3 bar gauge (barg) (30 kPa) to 6 barg (600 kPa). Thetemperature across reformate splitter 200 can be in the range from 50deg C and 300 deg C. It can be understood by one of skill in the artthat reformate splitter 200 can be designed to operate at a temperatureand pressure to produce the desired streams, such that different stagesin a single unit or different units when multiple used are deployed canoperate at different temperature and pressure from the other stages orunits. In at least one embodiment, where reformate splitter 200 is onedistillation column, the distillation column can include multiplesections in one vessel, where each section can have different operatingconditions.

Reformate splitter 200 can separate the components in dealkylationeffluent 6 to produce light gases stream 8, benzene stream 10, heavyhydrocarbons stream 12, C9 aromatic hydrocarbons stream 14, xylenestream 16, and toluene stream 20.

Light gases stream 8 can contain light gases. Light gases stream 8 canbe further processed to separate the light gases from hydrogen, whichcan be recycled to the dealkylation reactor. The light gases can befurther separated. Hydrocarbon gases containing one or two carbons canbe introduced to fuel gas blend. Hydrocarbon gases containing three orfour carbons can be introduced to an LPG gas plant. In an alternateembodiment, light gases stream 8 can be purged to atmosphere or can beburned to produce heat.

Benzene stream 10 can contain benzene. Benzene stream 10 can be stored,alternately can be further processed, alternately introduced totransalkylation reactor 600, and alternately can be disposed.

Heavy hydrocarbons 12 can contain C10+ aromatic hydrocarbons, includingC10+ aromatic hydrocarbons formed in dealkylation reactor 100 andunreacted C10+ aromatic hydrocarbons from reformate feed 2. Heavyhydrocarbons 12 can be purged from the system, alternately can befurther processed to recover hydrocarbons of value, and alternately canbe destroyed. Removing the C10+ aromatic hydrocarbons from the systemreduces accumulation of the inert C10+ aromatic hydrocarbons through therecycle streams from splitter column 700. Removing C10+ aromatichydrocarbons from the system can prolong the catalyst activity of thetransalkylation catalyst in transalkylation reactor 600 and can improvecatalyst performance.

Toluene stream 20 can contain toluene. A fraction of toluene stream 20can be introduced to toluene disproportionation reactor 300. In at leastone embodiment, toluene disproportionation reactor 300 is included whenthe amount of toluene in reformate feed 2 exceeds 150 wt % of the amountof trimethylbenzene. Toluene disproportionation reactor 300 can be anytype of unit capable of supporting a toluene disproportionationreaction, Reaction 1.

2CH₃C₆H₅↔C₆H₆+(CH₃)₂C₆H₄   (Reaction 1)

where CH₃C₆H₅ is toluene, C₆H₆ is benzene, and (CH₃)₂C₆H₄ is xylene. Inthe disproportionation reaction, a methyl group (—CH₃) is transferredfrom one toluene to another toluene to produce a benzene and a xylene.Toluene disproportionation reactor 300 can include a disproportionationcatalyst. The disproportionation catalyst can be any type of catalystcapable of catalyzing reaction 1. The disproportionation temperature intoluene disproportionation reactor 300 can be between 200 degreesCelsius (deg C) and 550 deg C. The disproportionation pressure intoluene disproportionation reactor 300 can be between 5 bar (500kilopascal (kPa)) and 50 bar (5000 kPa) and alternately between 10 bar(1000 kPa) and 50 bar (5000 kPa). The LHSV in toluene disproportionationreactor 300 can be between 1 per hour (hr⁻¹) and 20 hr⁻¹ and alternatelybetween 1 hr⁻¹ and 10 hr⁻¹. Reaction 1 is an equilibrium reaction with athermodynamic limit, where the reaction can be designed to balance theproduction of products with reaction conditions. The reaction productscan exit toluene disproportionation reactor 300 as disproportionationeffluent 30. Disproportionation effluent 30 can include toluene,benzene, mixed xylenes, and combinations of the same. Disproportionationeffluent 30 can be introduced to splitter column 700.

In at least one embodiment, toluene split stream 22 can be separatedfrom toluene stream 20 and introduced to transalkylation reactor 600.Toluene split stream 22 can have a flow rate to maintain a ratio oftoluene to trimethylbenzene in transalkylation reactor 600 in the rangeof 0.3 to 3. In at least one embodiment, when the flow rate of toluenestream 20 is insufficient to maintain a ratio of toluene totrimethylbenzene in the range of 0.3 to 3, an alternate stream oftoluene can be introduced to transalkylation reactor 600. Maintainingthe ratio of toluene to trimethylbenzene in the range of 0.3 to 3increases the production of xylene, alternately in the range of 0.5 to1.5, and alternately in the range of 0.8 to 1.2. Maintaining the ratioof toluene to trimethylbenzene close to 1 results in greater selectivityof xylene in the product.

In at least one embodiment, as shown in FIG. 2, the process to maximizexylene production is in the absence of toluene disproportionationreactor 300. In a process in the absence of toluene disproportionationreactor 300 any excess toluene not separated in toluene split stream 22can be introduced to splitter column 700. In at least one embodiment,the entire flow of toluene stream 20 is introduced to transalkylationreactor 600 through toluene split stream 22.

Xylene stream 16 can contain mixed xylenes. Xylene stream 16 can beintroduced to p-xylene separation unit 400. P-xylene separation unit 400can be any type of separation unit capable of separating p-xylene from astream containing mixed xylenes. Examples of p-xylene separation unit400 can include a crystallization unit or an adsorption unit. P-xylenesare separated from xylene stream 16 in p-xylene separation unit 400 toproduce p-xylene depleted stream 40 and p-xylene product 42. P-xyleneproduct 42 can contain p-xylene. P-xylene product 42 can be stored,alternately can be further processed, and alternately can be used asfinished product stream. P-xylene depleted stream 40 can containm-xylene and o-xylene. P-xylene depleted stream 40 can be introduced toisomerization unit 500.

P-xylene separation unit 400 cannot be used to separate components otherthan p-xylene; one of skill in the art will appreciate that a p-xylene.Advantageously, separating C9 aromatics stream 14 upstream of p-xyleneseparation unit 400 increases the effectiveness of separating p-xylenefrom xylene stream 16 in p-xylene separation unit 400.

Isomerization unit 500 can be any type of unit capable of supportingxylene isomerization reactions, including associated units andinstrumentation. Isomerization unit can include an isomerizationcatalyst. The isomerization catalyst can be any type of catalyst capableof catalyzing the isomerization reaction of m-xylene and o-xylene top-xylene. Examples of the isomerization catalyst can include a surfacemodified zeolite catalyst such as a zeolite beta, an HZSM-5, an MCM-49,and ZSM-12. The isomerization temperature in isomerization unit 500 canbe between 200 degrees Celsius (deg C) and 550 deg C. The isomerizationpressure in isomerization unit 500 can be between 5 bar (500 kilopascal(kPa)) and 50 bar (5000 kPa) and alternately between 10 bar and 50 bar.The LHSV in isomerization unit 500 can be between 1 per hour (hr⁻¹) and20 hr⁻¹ and alternately between 1 hr⁻¹ and 10 hr⁻¹. The mixed xylenesproduced in isomerization unit 500 can exit as isomerization effluent50. Isomerization effluent 50 can be introduced to splitter column 700.

C9 aromatics stream 14 can contain C9 aromatic hydrocarbons, includingC9 aromatic hydrocarbons formed in dealkylation reactor 100 andunreacted C9 aromatic hydrocarbons from reformate feed 2. In at leastone embodiment, C9 aromatics stream 14 is in the absence of C10+aromatic hydrocarbons. In at least one embodiment, C9 aromatics stream14 contains less than 5 wt % C10+ aromatic hydrocarbon. C9 aromaticsstream 14 along with hydrogen stream 5 can be introduced totransalkylation reactor 600.

Hydrogen stream 5 can be any stream containing hydrogen gas. Hydrogenstream 5 can be a stream of pure hydrogen from a fresh hydrogen source.In at least one embodiment, hydrogen gas can be recovered from anotherpart of the process and can be recycled as hydrogen stream 5 andintroduced to transalkylation reactor 600. Hydrogen stream 5 can be froma hydrogen source in a refinery and can contain light hydrocarbons.

Transalkylation reactor 600 can be a fixed bed reactor or a fluidizedbed reactor. The transalkylation temperature in transalkylation reactor600 can be between 200 deg C and 550 deg C. The transalkylation pressurein transalkylation reactor 600 can be between 5 bar (500 kPa) and 50 bar(5000 kPa), alternately 10 bar (1000 kPa) and 50 bar (5000 kPa), andalternately 10 bar (1000 kPa) and 40 bar (4000 kPa). The LHSV can bebetween 0.2 hr⁻¹ and 20 hr⁻¹ and alternately between 0.5 hr⁻¹ and 6hr⁻¹. The volumetric flow ratio of hydrogen gas to hydrocarbons can bein the range of 0 to 8. The operating conditions can be set to maximizethe production of xylenes. The temperature can have a greater influenceon the transalkylation reaction than pressure. It is understood thathigher temperatures, higher pressures, and lower LHSV favortransalkylation reactions, while higher temperatures can lead tocatalyst deactivation and therefore, the operating conditions must bebalanced to maximize production and minimize catalyst deactivation. Dueto the conversion of C9+ aromatic hydrocarbons in dealkylation reactor100, the loading on transalkylation reactor can be reduced.

Transalkylation reactor 600 can include a transalkylation catalyst. Thetransalkylation catalyst can include any catalyst capable of catalyzingtransalkylation reactions. Examples of transalkylation catalysts includebifunctional catalysts as described in U.S. Pat. No. 5,866,741. Thetransalkylation catalyst can be selected to selectively convert one ormore of the C9+ aromatic hydrocarbons over the others in transalkylationreactions. In at least one embodiment, the transalkylation catalyst canbe selected to selectively convert trimethylbenzenes to mixed xylenes.Advantageously, using a catalyst designed to catalyze transalkylationreactions increases the efficiency and performance as compared to acatalyst designed to perform transalkylation, disproportionation andisomerization. Using a catalyst targeted to transalkylation reactionsincreases the production of mixed xylenes.

There are two primary reactions that occur in transalkylation reactor600 to form mixed xylenes. One reaction is Reaction 1 described here.The second reaction is an equilibrium transalkylation reaction oftoluene and trimethylbenzene:

CH₃C₆H₅+(CH₃)₃C₆H₃↔2(CH₃)₂C₆H₄   (Reaction 2)

where CH₃C₆H₅ is toluene, (CH₃)₃C₆H₃ is trimethylbenzene, and2(CH₃)₂C₆H₄ is xylene. In addition to the two primary reactions, othertransalkylation reactions can convert C9+ aromatic hydrocarbons totoluene, benzene, mixed xylenes, and light gases. The reaction productsproduced in transalkylation reactor 600 can exit as transalkylationeffluent 60. Transalkylation effluent 60 can contain mixed xylenes,toluene, benzene, light gases, and C9+ aromatic hydrocarbons. The C9+aromatic hydrocarbons in transalkylation effluent 60 can include C9+aromatic hydrocarbons formed in transalkylation reactor, C9+ aromatichydrocarbons formed in the dealkylation reactor, and C9+ aromatichydrocarbons from reformate feed 2. Transalkylation effluent 60 can beintroduced to splitter column 700.

Splitter column 700 can be any type of separation unit capable ofseparating a stream into its components. Splitter column 700 can be anytype of separation unit capable of separating a stream into itscomponent parts. In at least one embodiment, splitter column 700 can beone splitter column designed to separate the feed stream into multiplesplit streams. In at least one embodiment, splitter column 700 can bemultiple splitter columns in series designed to separate one componentfrom the feed stream. In at least one embodiment, splitter column 700can be one or more distillation units. The pressure across splittercolumn 700 can be in the range from 0.3 bar gauge (barg) (30 kPa) to 6barg (600 kPa). The temperature across splitter column 700 can be in therange from 50 deg C and 300 deg C. It can be understood by one of skillin the art that splitter column 700 can be designed to operate at atemperature and pressure to produce the desired streams, such thatdifferent stages in a single unit or different units when multiple usedare deployed can operate at different temperature and pressure from theother stages or units. In at least one embodiment, where splitter column700 is one distillation column, the distillation column can includemultiple sections in one vessel, where each section can have differentoperating conditions.

Splitter column 700 can separate the components in disproportionationeffluent 30, isomerization effluent 50, and transalkylation effluent 60to produce separated light gases 70, benzene recycle 72, toluene recycle74, xylene recycle 76, and C9+ aromatics recycle 78. Separated lightgases 70 can be processed further or stored. Separated light gases 70can be processed further or stored. Benzene recycle 72, toluene recycle74, and C9+ aromatics recycle 78 can be recycled to transalkylationreactor 600. Xylene recycle 76 can be recycled to p-xylene separationunit 400. Adding benzene from benzene recycle 72 to transalkylationreactor 600 can drive the equilibrium toward of benzene through Reaction1, which reduces the consumption of toluene in Reaction 1. By reducingthe consumption of toluene, more toluene is available for production ofmixed xylenes through Reaction 2. The amount of benzene added throughbenzene recycle 72 can be in the range between 1 wt % and 10 wt %. Theamount of benzene added through benzene recycle 72 can be a function ofthe transalkylation temperature in transalkylation reactor 600, andtherefore, the transalkylation temperature can be monitored withtemperature sensors or other instrumentation, such that the flow rate ofbenzene recycle 72 can be adjusted across the run in transalkylationreactor 600. In at least one embodiment, the transalkylation temperatureis different at the beginning of the run and at the end of the run. Inat least one embodiment, benzene stream 10 can be introduced to 600transalkylation reactor to maintain the thermodynamic equilibrium.

C9+ aromatics recycle 78 can contain an amount of C10 hydrocarbons thatdoes not participate in any of the reactions in transalkylation reactor600. In at least one embodiment, C9+ aromatics recycle 78 can includetrace amounts of C11 hydrocarbons and trace amounts of C12 hydrocarbons.A slip stream from C9+ aromatics recycle 78 can be purged from thesystem to prevent build up in the loop between transalkylation reactor600 and splitter column 700.

The yield of p-xylene in p-xylene product 42 is in the range between 75wt % and 82 wt %.

Referring to FIG. 3, an alternate embodiment of the process to maximizeproduction of xylenes is provided with reference to FIG. 1. Mixed lightaromatics stream 18 can be separated in reformate splitter 200. Mixedlight aromatics stream 18 can include toluene, benzene, non-aromaticsand combinations of the same. Mixed light aromatics stream 18 canfurther include non-aromatics. Mixed light aromatics stream 18 can beintroduced to aromatic extraction unit 800.

Aromatic extraction unit 800 can be any type of extraction unit capableof separating aromatic components from other components in a stream.Examples of aromatic extraction unit can include liquid-liquidextraction and extractive distillation, both of which use solvents.Examples of solvents suitable for use in aromatic extraction unit 800can include sulfolane (C₄H₈O₂S), furfural (C₅H₄O₂), tetraethylene glycol(C₈H₁₈O₅), dimethylsulfoxide (C₂H₆OS), and N-methyl-2-pyrrolidone(C₅H₉NO). The particular conditions in aromatic extraction unit 800depend on the type of extraction technology selected. Aromaticextraction unit 800 can be included in embodiments where theconcentration of non-aromatics is at a level that can disrupt thefunction or performance of the units of the system.

Non-aromatic hydrocarbons in mixed light aromatics stream 18 can beseparated in aromatic extraction unit 800 to produce non-aromaticsraffinate stream 82 and aromatics extract stream 80. Aromatics extractstream 80 can include toluene, benzene, and combinations of the same.Non-aromatics raffinate stream 82 can include non-aromatics. Removingnon-aromatics can remove the non-aromatics from the downstream units andreduce dilution of the reactants in the downstream units, includingtransalkylation reactor 600 and p-xylene separation unit 400.Advantageously, aromatic extraction unit 800 separates non-aromaticsthat are not separable from C6 and C7 aromatics in reformate splitter200. Non-aromatic hydrocarbons can include paraffins, iso-paraffins,naphthenes, and combinations of the same.

In at least one embodiment of the process according to FIG. 3, a portionof benzene recycle 72 can be purged from the system, while the remainderis introduced to transalkylation reactor 600. As noted above withrespect to FIG. 1, the amount of benzene added through benzene recycle72 can be adjusted based on the reaction conditions in transalkylationreactor 600 to optimize the production of mixed xylenes. Purging aportion of benzene recycle 72 provides a method for controlling theamount of benzene added through benzene recycle 72, avoids accumulatingexcess benzene in the system, and prevents dilution of the reactants intransalkylation reactor 600. A portion of benzene recycle 72 can bepurged when benzene recycle 72 contains more benzene that is required todrive the equilibrium reactions in transalkylation reactor 600.

Referring to FIG. 4, an alternate embodiment of the process to maximizeproduction of xylenes is provided with reference to FIG. 1 and FIG. 3.Dealkylation effluent 6 is introduced to aromatic extraction unit 800.Non-aromatics in dealkylation effluent 6 can be separated in aromaticextraction unit 800 to produce aromatic feed 84 and raffinate effluent86. Raffinate effluent 86 can include the non-aromatics separated fromdealkylation effluent 6, including light hydrocarbons. Aromatic feed 84can include C6 to C9 aromatic hydrocarbons produced in dealkylationeffluent 100 and aromatic hydrocarbons present in reformate feed 2.

Aromatic feed 84 can be introduced to reformate splitter 200 to producebenzene stream 10, heavy hydrocarbons 12, C9 aromatic hydrocarbonsstream 14, xylene stream 16, and toluene stream 20. Toluene stream 20,in whole, can be introduced to transalkylation reactor 600.

Referring to FIG. 6, an alternate embodiment of the process to maximizeproduction of xylenes is provided with reference to FIG. 1. Reformatefeed 2 is introduced to reformate splitter 200. Reformate splitter 200can separate the components in reformate feed 2 to produce light gasesstream 8, benzene stream 10, combined heavy stream 64, xylene stream 16,and toluene stream 20. Separating the reformate stream can reduce theload in dealkylation reactor 100. Benzene and toluene do not react indealkylation reactor 100 and therefore, removing them from the reactorenhances catalyst performance by reducing dilution. Combined heavystream 64 can include the C9+ aromatic hydrocarbons present in reformatefeed 2 and the trace amount of C9+ aromatic hydrocarbons produced intransalkylation reactor 600.

Combined heavy stream 64 can be introduced to dealkylation reactor 100.Hydrogen feed 4 can be introduced to dealkylation reactor 100.Dealkylation reactor 100 operates as described with reference to FIG. 1.The inclusion of dealkylation reactor 100 can increase the amount ofxylenes produced by increasing conversion of C10+ aromatic hydrocarbons.

Dealkylation effluent 6 can be introduced to dealkylation splitter 900.Dealkylation splitter 900 can separate the components in dealkylationeffluent 6 to produce light recycle stream 94, C9 stream 96, and C10+aromatics stream 90. Light recycle stream 94 can contain light gases,benzene, toluene, xylenes, and combinations of the same. Light recyclestream 94 can be recycled to reformate splitter 200. C10+ aromaticsstream 90 can contain C10+ aromatic hydrocarbons. C10+ stream 90 can beintroduced to transalkylation reactor 600. Introducing C10+ stream 90 totransalkylation reactor 600 can maintain the desired thermodynamics intransalkylation reactor 600. A purge stream can be separated from C10+stream 90 as purge stream 92. The flow rate of purge stream 92 can bedetermined based on the overall volumetric flow rate of C10+ stream 90.The flow rate of purge stream 92 can be optimized to reduce loading ofC10+ aromatic hydrocarbons in transalkylation reactor 600. Reducedloading can increase the conversion of MEB to toluene in transalkylationreactor 600. In at least one embodiment, the conversion rate of MEB intransalkylation reactor 600 is equal to or greater than 98%.Additionally, reduced loading can reduce fouling of the transalkylationcatalyst in transalkylation reactor 600 prolonging the run time oftransalkylation reactor 600 and can reduce the size of transalkylationreactor 600. In at least one embodiment, the entire volumetric flow rateof C10+ stream 90 flows out through purge stream 92. In at least oneembodiment, the entire volumetric flow rate of C10+ stream 90 isintroduced to transalkylation reactor 600.

Optionally, in at least one embodiment, purge stream 92 can be recycledto dealkylation reactor 100. The hydrogen pressure and severity of theoperation conditions can be adjusted to achieve full conversion of theC10+ aromatic hydrocarbons to C9 aromatic hydrocarbons and compoundswith less than 9 carbon atoms.

C10+ stream 90 can be introduced to transalkylation reactor 600 alongwith C9 stream 96, benzene stream 10, and toluene split stream 22. Asdescribed with reference to FIG. 1, hydrogen stream 5 can be introducedto transalkylation reactor 600. In the embodiment described withreference to FIG. 6, the entire volumetric flow of benzene stream 10 isintroduced to transalkylation reactor 600. Benzene stream 10 can beintroduced to transalkylation reactor 600 to maintain the thermodynamicequilibrium, which drives conversion of toluene to xylenes. The presenceof benzene can drive equilibrium toward other reaction products, such asxylenes. Optionally, in certain embodiments, a portion of benzene streamcan be separated as benzene split stream 62. The volumetric flow ofbenzene split stream 62 can be determined by the improved xylene yield.The amount of benzene introduced to the transalkylation reactor throughbenzene stream 10 can be in the range between 0.01 wt % and 5 wt % withthe remaining amount removed from the system through benzene splitstream 62.

In the embodiment described with reference to FIG. 6, the entirevolumetric flow of toluene stream 20 is introduced to transalkylationreactor 600 through toluene split stream 22. Optionally, in certainembodiments, a portion of toluene stream 20 can be introduced to toluenedisproportionation reactor 300. Toluene disproportionation reactor 300can be included when the amount of toluene in reformate feed 2 exceeds150 wt % of the amount of trimethylbenzene. In the embodiments thatinclude toluene disproportionation reactor 300, the volumetric flow rateof toluene split stream 22 can be adjusted to maintain a ratio oftoluene to trimethylbenzene in the range of 0.3 to 3 in transalkylationreactor 600. Toluene disproportionation reactor 300 operates asdescribed with reference to FIG. 1. In embodiments where toluenedisproportionation reactor 300 is present, disproportionation effluent30 can be recycled to reformate splitter 200.

Transalkylation reactor 600 operates as described with reference toFIG. 1. Transalkylation effluent 60 can be introduced to reformatesplitter 200. Isomerization effluent 50, produced as described withreference to FIG. 1, can be recycled to reformate splitter 200.

Using reformate splitter 200 to separate components indisproportionation effluent 30, isomerization effluent 50, andtransalkylation effluent 60 can reduce complexity of the system. In analternate embodiment, disproportion effluent 30, isomerization effluent50, and transalkylation effluent 60 can be introduced to splitter column700 as described with respect to FIG. 1.

Advantageously, the position of the dealkylation reactor upstream of thetransalkylation produces toluene not present in the heavy reformatefeed, toluene is a reactant in transalkylation reactions to producexylene, thus a process with the dealkylation reactor upstream of thetransalkylation increases xylene production. Advantageously, theposition of the dealkylation reactor upstream of the transalkylationreactor reduces the amount of C9 aromatic hydrocarbons and C10+ aromatichydrocarbons being introduced to the transalkylation reactor.

Advantageously, the use of both a reformate splitter and a splittercolumn increase the ability to control the recycle of components to thetransalkylation reactor. Advantageously, the use of both a reformatesplitter and a splitter column means that the splitter column handlesonly aromatic hydrocarbons, while the reformate splitter handles boththe aromatic hydrocarbons and non-aromatic hydrocarbons.

Both the dealkylation reactor and the transalkylation reactor are in theabsence of methanol and in the absence of methylation reactions, whichare irreversible reactions that add a methyl group to a compound. In atleast one embodiment, the heavy reformate feed is in the absence ofethylbenzene. The process to maximize xylene production is in theabsence of a debutanizer. The process to maximize xylene production isin the absence of a depheptanizer. In at least one embodiment, thesystem for producing p-xylene is in the absence of a hydrogenationreaction or hydrogenation reactor, where a hydrogenation reactionreactor refers to a reaction that generally adds hydrogen to ahydrocarbon for the purposes of saturating the hydrocarbon. In at leastone embodiment, the system is in the absence of a hydrodealkylationreaction or a hydrodealkylation reactor, where a hydrodealkylationreaction is a reaction that involves the detachment of an alkyl groupfrom an aromatic ring.

Dealkylation reactor 100 is not a reforming unit. A reforming unitconverts naphthas into an aromatic hydrocarbons-rich product over aplatinum-containing catalyst at high temperatures and hydrogen pressure.A reforming unit has the goal of increasing the octane number of thehydrocarbon feed stream to the reforming unit. The primary reactions ina reforming unit are dehydrogenation reactions and cracking reactions.

EXAMPLES

The following examples were carried out using a simulation program.

Example 1 was simulated as the process for maximizing xylene productionaccording to FIG. 1. The process was simulated using ASPEN. The processfor maximizing xylene production includes a dealkylation reactor, atransalkylation reactor, p-xylene separation and xylene isomerization.

TABLE 1 Composition and flow rates for Example 1 2 4 6 8 10 12 14 16 2022 30 40 42 Mass Flow 1598 10 1608 86 73 73 506 497 373 369 4 4544 1320Hydrogen kg/hr 0 10 7 7 0 0 0 0 0 0 0 0 0 Light Gas kg/hr 0 0 79 79 0 00 0 0 0 0 0 0 Benzene kg/hr 49 0 73 0 73 0 0 0 0 0 2 0 0 Toluene kg/hr213 0 373 0 0 0 0 0 373 369 0 0 0 p-xylene kg/hr 97 0 119 0 0 0 0 119 00 1 0 1320 o-xylene & kg/hr 291 0 378 0 0 0 0 378 0 0 1 4544 0 m-xyleneMEB kg/hr 211 0 3 0 0 0 3 0 0 0 0 0 0 TMB kg/hr 565 0 503 0 0 0 503 0 00 0 0 0 C10+ kg/hr 171 0 73 0 0 73 0 0 0 0 0 0 0 50 5 60 70 72 74 76 78Mass Flow 4544 100 2265 156 67 649 5367 573 Hydrogen kg/hr 0 100 96 96 00 0 0 Light Gas kg/hr 0 0 61 61 0 0 0 0 Benzene kg/hr 0 0 65 0 67 0 0 0Toluene kg/hr 0 0 649 0 0 649 0 0 p-xylene kg/hr 1045 0 156 0 0 0 1201 0o-xylene & kg/hr 3499 0 666 0 0 0 4166 0 m-xylene MEB kg/hr 0 0 3 0 0 00 3 TMB kg/hr 0 0 503 0 0 0 0 503 C10+ kg/hr 0 0 67 0 0 0 0 67

Example 2

Example 2 is a comparative example in the absence of a dealkylationreactor. The process was simulated according to the process of FIG. 5.FIG. 5 depicts a process in the absence of dealkylation reactor 100,where reformate feed 2 is introduced directly to reformate splitter 200.

TABLE 2 Composition and flow rates for Example 2. 2 10 12 14 16 20 22 3040 42 Mass Flow 1598 49 146 802 388 213 211 2 4439 1290 Hydrogen kg/hr 00 0 0 0 0 0 0 0 0 Light Gas kg/hr 0 0 0 0 0 0 0 0 0 0 Benzene kg/hr 4949 0 0 0 0 0 1 0 0 Toluene kg/hr 213 0 0 0 0 213 211 0 0 0 p-xylenekg/hr 97 0 0 0 97 0 0 0 0 1290 o-xylene & kg/hr 291 0 0 0 291 0 0 1 44390 m-xylene MEB kg/hr 211 0 0 211 0 0 0 0 0 0 TMB kg/hr 565 0 0 565 0 0 00 0 0 C10+ kg/hr 171 0 146 25 0 0 0 0 0 0 50 5 60 70 72 74 76 78 MassFlow 4439 100 2486 214 73 449 5342 851 Hydrogen kg/hr 0 100 93 93 0 0 00 Light Gas kg/hr 0 0 120 120 0 0 0 0 Benzene kg/hr 0 0 72 0 73 0 0 0Toluene kg/hr 0 0 449 0 0 449 0 0 p-xylene kg/hr 1021 0 172 0 0 0 1193 0o-xylene kg/hr 3418 0 729 0 0 0 4148 0 & m-xylene MEB kg/hr 0 0 211 0 00 0 211 TMB kg/hr 0 0 565 0 0 0 0 565 C10+ kg/hr 0 0 74 0 0 0 0 74

Comparing the results of Example 1 and Example 2 shows that the additionof dealkylation reactor 100 reduces the flow rate of C9 aromatichydrocarbons introduced to transalkylation reactor by 9.1%. In part thereduced flow rate is due to the conversion of methylethylbenzene totoluene in dealkylation reactor 100 in the process to maximize xylenesas embodied in FIG. 1. In addition, the results indicate that theaddition of dealkylation reactor 100 increases the conversion of C10+aromatic hydrocarbons to C6 to C8 aromatic hydrocarbons, where thenon-converted compounds can be purged from the system as part of heavyhydrocarbons 12.

Example 3

Example 3 is an example of a simulated process described with referenceto FIG. 6. FIG. 6. includes dealkylation reactor 100 downstream fromreformate splitter 200.

TABLE 3 Composition and flow rates for Example 3. 2 8 10 62 20 22 16 646 90 92 94 Mass Flow 1598 140 148 74 932 923 5884 1462 1470 103 76 458Hydrogen kg/hr 0 7 0 0 0 0 0 0 3 0 0 3 Light Gas kg/hr 0 133 0 0 0 0 0 059 0 0 59 Benzene kg/hr 49 0 148 74 0 0 0 0 37 0 0 37 Toluene kg/hr 2130 0 0 932 923 0 0 161 0 0 161 p-xylene kg/hr 97 0 0 0 0 0 1325 0 39 0 039 o-xylene & kg/hr 291 0 0 0 0 0 4558 0 159 0 0 159 m-xylene MEB kg/hr211 0 0 0 0 0 0 212 3 0 0 0 TMB kg/hr 565 0 0 0 0 0 0 1019 907 0 0 0C10+ kg/hr 171 0 0 0 0 0 0 231 103 103 76 0 96 4 5 42 40 50 30 60 MassFlow 909 8 8 1325 4558 4558 9 1942 Hydrogen kg/hr 0 8 8 0 0 0 0 4 LightGas kg/hr 0 0 0 0 0 0 0 73 Benzene kg/hr 0 0 0 0 0 0 4 58 Toluene kg/hr0 0 0 0 0 0 0 558 p-xylene kg/hr 0 0 0 1325 0 1048 2 139 o-xylene &kg/hr 0 0 0 0 4558 3510 4 595 m-xylene MEB kg/hr 3 0 0 0 0 0 1 TMB kg/hr907 0 0 0 0 0 0 453 C10+ kg/hr 0 0 0 0 0 0 0 60

Example 4

Example 4 is a comparative example. The process was simulated accordingto the process of FIG. 7. The process of FIG. 7 can be understood withreference to FIG. 6, where the process is in the absence of dealkylationreactor 100, such that no dealkylation occurs upstream oftransalkylation reactor 600. Combined heavy stream 64 can flow directlyinto dealkylation splitter 900.

TABLE 4 Composition and flow rates for Example 4. 2 8 10 62 20 22 16 6490 92 Mass Flow 1598 122 126 63 702 695 5866 1800 22 101 Hydrogen kg/hr0 1 0 0 0 0 0 0 0 0 Light Gas kg/hr 0 120 0 0 0 0 0 0 0 0 Benzene kg/hr49 0 126 63 0 0 0 0 0 0 Toluene kg/hr 213 0 0 0 702 695 0 0 0 0 p-xylenekg/hr 97 0 0 0 0 0 1320 0 0 0 o-xylene & kg/hr 291 0 0 0 0 0 4546 0 0 0m-xylene MEB kg/hr 211 0 0 0 0 0 0 422 0 0 TMB kg/hr 565 0 0 0 0 0 01131 0 0 C10+ kg/hr 171 0 0 0 0 0 0 248 22 101 96 4 5 42 40 50 30 60Mass Flow 1578 8 8 1320 4546 4546 7 2466 Hydrogen kg/hr 0 8 8 0 0 0 0 1Light Gas kg/hr 0 0 0 0 0 0 0 120 Benzene kg/hr 0 0 0 0 0 0 3 74 Toluenekg/hr 0 0 0 0 0 0 0 489 p-xylene kg/hr 0 0 0 1320 0 1046 1 177 o-xylene& kg/hr 0 0 0 0 4546 3500 3 752 m-xylene MEB kg/hr 422 0 0 0 0 0 211 TMBkg/hr 1131 0 0 0 0 0 0 565 C10+ kg/hr 25 0 0 0 0 0 0 76

Comparing the results of Examples 3 and Example 4 shows that theaddition of dealkylation reactor 100 reduces the flow rate of C9aromatic hydrocarbons introduced to transalkylation reactor 600.

Although the embodiments have been described in detail, it should beunderstood that various changes, substitutions, and alterations can bemade without departing from the principle and scope. Accordingly, thescope of the present embodiments should be determined by the followingclaims and their appropriate legal equivalents.

There various elements described can be used in combination with allother elements described here unless otherwise indicated.

The singular forms “a”, “an” and “the” include plural referents, unlessthe context clearly dictates otherwise.

Optional or optionally means that the subsequently described event orcircumstances may or may not occur. The description includes instanceswhere the event or circumstance occurs and instances where it does notoccur.

Ranges may be expressed here as from about one particular value to aboutanother particular value and are inclusive unless otherwise indicated.When such a range is expressed, it is to be understood that anotherembodiment is from the one particular value to the other particularvalue, along with all combinations within said range.

Throughout this application, where patents or publications arereferenced, the disclosures of these references in their entireties areintended to be incorporated by reference into this application, in orderto more fully describe the state of the art to which the inventionpertains, except when these references contradict the statements madehere.

As used here and in the appended claims, the words “comprise,” “has,”and “include” and all grammatical variations thereof are each intendedto have an open, non-limiting meaning that does not exclude additionalelements or steps.

That which is claimed is:
 1. A system for producing para-xylene(p-xylene) from a reformate feed, the system comprising: a reformatesplitter, the reformate splitter configured to separate the reformatefeed, wherein the reformate feed comprises aromatic hydrocarbons,wherein the aromatic hydrocarbons are selected from the group consistingof benzene, toluene, mixed xylenes, carbon-nine plus (C9+) aromatichydrocarbons, and combinations of the same, wherein the reformatesplitter produces a light gases stream, a combined heavy stream, abenzene stream, a xylene stream, and toluene stream, wherein the C9aromatics stream comprises C9 aromatic hydrocarbons, and wherein thexylene stream comprises mixed xylenes, wherein the mixed xylenescomprises p-xylenes, and wherein the mixed light aromatics streamcomprises toluene; a dealkylation reactor fluidly connected to thereformate splitter, the dealkylation reactor configured to convert theC9+ aromatic hydrocarbons to carbon-six (C6) to carbon-eight (C8)aromatic hydrocarbons in the presence of a dealkylation catalyst toproduce a dealkylation effluent, wherein the dealkylation reactor is ata dealkylation temperature, wherein the dealkylation reactor is at adealkylation pressure, wherein the dealkylation effluent comprisesaromatic hydrocarbons such that an amount of C9+ aromatic hydrocarbonsin the dealkylation effluent is less than the amount of C9+ aromatichydrocarbons in the reformate feed; a dealkylation splitter fluidlyconnected to the dealkylation reactor, the dealkylation splitterconfigured to separate the dealkylation effluent to produce a lightrecycle stream, C10+ aromatics stream, and C9 stream; a transalkylationreactor fluidly connected to the dealkylation reactor and fluidlyconnected to the reformate splitter, the transalkylation reactorconfigured to convert toluene and C9 aromatic hydrocarbons in thepresence of a transalkylation catalyst to produce a transalkylationeffluent, wherein the transalkylation reactor is at a transalkylationtemperature, wherein the transalkylation reactor is at a transalkylationpressure, wherein the toluene is supplied by a toluene split stream,wherein the toluene split stream is separated from the toluene stream,wherein the transalkylation catalyst is operable to catalyzetransalkylation reactions, wherein the transalkylation effluentcomprises C6 to C9+ aromatic hydrocarbons; a p-xylene separation unitfluidly connected to a reformate splitter, the p-xylene separation unitconfigured to separate p-xylenes from the xylene stream to produce ap-xylene product and a p-xylene depleted stream, wherein the p-xyleneproduct comprises p-xylenes, wherein the p-xylene depleted streamcomprises meta-xylene (m-xylene) and ortho-xylene (o-xylene); and anisomerization unit fluidly connected to the p-xylene separation unit,wherein the isomerization unit is configured to convert the m-xylene ando-xylene in the p-xylene depleted stream to produce an isomerizationeffluent, wherein the isomerization unit is at an isomerizationtemperature, wherein the isomerization unit is at an isomerizationpressure, wherein the isomerization effluent comprises C8 aromatichydrocarbons.
 2. The system of claim 1, wherein a flow rate of thetoluene split stream is operable to maintain a ratio of toluene totrimethylbenzene in the transalkylation reactor in a range of 0.3 to 3.3. The system of claim 1, wherein the dealkylation temperature isbetween 200° C. and 550° C., further wherein the dealkylation pressureis between 5 bar and 50 bar, and wherein a liquid hourly space velocityin the dealkylation reactor is between 1 hr⁻¹ and 20 hr⁻¹.
 4. The systemof claim 1, wherein the transalkylation temperature is between 200° C.and 550° C., further wherein the transalkylation pressure is between 10bar and 50 bar, and wherein a liquid hourly space velocity in thetransalkylation reactor is between 0.2 hr⁻¹ and 20 hr⁻¹.
 5. The systemof claim 1, wherein the isomerization temperature is between 200° C. and550° C., further wherein the isomerization pressure is between 5 bar and50 bar, and wherein a liquid hourly space velocity in the isomerizationunit is between 1 hr⁻¹ and 20 hr⁻¹.
 6. The system of claim 1, furthercomprising a toluene disproportionation reactor fluidly connected to thereformate splitter, the toluene disproportionation reactor configured toconvert a fraction of the toluene from the toluene stream to produce adisproportionation effluent, wherein the disproportionation effluentcomprises C6 to C8 aromatic hydrocarbons.